Slurry phase polymerisation process

ABSTRACT

A process comprising polymerising in a loop reactor of continuous tubular construction an olefin monomer optionally together with an olefin comonomer in the presence of a polymerisation catalyst in a diluent to produce a slurry comprising solid particulate olefin polymer and the diluent wherein the internal diameter of at least 50% of the total length of the reactor is at least 700 millimeters and the solids concentration in the reactor is at least 20 volume % is disclosed.

BACKGROUND OF THE INVENTION

The present invention is concerned with olefin polymerisation in slurryphase loop reactors.

Slurry phase polymerisation of olefins is well known wherein an olefinmonomer and optionally olefin comonomer are polymerised in the presenceof a catalyst in a diluent in which the solid polymer product issuspended and transported.

DESCRIPTION OF THE INVENTION

This invention is specifically related to polymerisation in a loopreactor where the slurry is circulated in the reactor typically by meansof a pump or agitator. Liquid full loop reactors are particularly wellknown in the art and are described for example in U.S. Pat. Nos.3,152,872, 3,242,150 and 4,613,484.

Polymerisation is typically carried out at temperatures in the range50-125° C. and at pressures in the range 1-100 bara. The catalyst usedcan be any catalyst typically used for olefin polymerisation such aschromium oxide, Ziegler-Natta or metallocene-type catalysts. The productslurry comprising polymer and diluent, and in most cases catalyst,olefin monomer and comonomer can be discharged intermittently orcontinuously, optionally using concentrating devices such ashydrocyclones or settling legs to minimise the quantity of fluidswithdrawn with the polymer.

The loop reactor is of a continuous tubular construction comprising atleast two, for example four, vertical sections and at least two, forexample four horizontal sections. The heat of polymerisation istypically removed using indirect exchange with a cooling medium,preferably water, in jackets surrounding at least part of the tubularloop reactor. The volume of the loop reactor can vary but is typicallyin the range 20 to 120 m³. the loop reactors of the present inventionare of this generic type.

Maximum commercial scale plant capacities have increased steadily overthe years. Growing operating experience over the last few decades hasled to operation of increasingly high slurry and monomer concentrationsin reaction loops. The increase in slurry concentrations has typicallybeen achieved with increased circulation velocities achieved for exampleby higher reactor circulation pump head or multiple circulation pumps asillustrated by EP 432555 and EP 891990. The increase in solids loadingis desirable to increase reactor residence time for a fixed reactorvolume and also to reduce downstream diluent treatment and recyclingrequirements. The increased velocity and head requirement of the loophas however led to increasing pump design sizes and complexity, andenergy consumptions as slurry concentrations increase. This has bothcapital and operating cost implications.

Historically the circulation velocity in the reaction loop has typicallybeen maximised to ensure maintenance of good thermal, compositional andparticle distribution across the reactor cross-section, particularly theavoidance of solids settling, stable flow characteristics, or excessivesolids concentrations at the pipe wall rather than reduced to minimisepressure drop/power in the polymerisation loop.

Inadequate cross-sectional distribution could lead to increased fouling,reduced heat transfer and reduced polymer productivity and homogeneity.Construction and commissioning of new commercial plants is veryexpensive and therefore new designs seek to avoid or minimise changes tooperating parameters that are seen to increase risk to the successfuloperation of the new unit.

In accordance with the present invention there is provided a processcomprising polymerising in a loop reactor of a continuous tubularconstruction an olefin monomer optionally together with an olefincomonomer in the presence of a polymerisation catalyst in a diluent toproduce a slurry comprising solid particulate olefin polymer and thediluent wherein the average internal diameter of at least 50% of thetotal length of the reactor is at least 700 millimeters and the solidsconcentration in the reactor is at least 20 volume %

One advantage of the present invention is that the specific energyconsumption of the reactor (i.e. the energy consumed per unit weight ofpolymer produced) is reduced whilst maintaining a given reactorresidence time and avoiding unacceptable reactor fouling. The inventionis especially advantageous when it is desired to design and operate aplant at high solids loadings when it has previously been considerednecessary to use what have now been found to be excessively high loopcirculation velocities.

This invention relates to a method and apparatus for continuouspolymerization of olefins, preferably alpha mono olefins, in anelongated tubular closed loop reaction zone. The olefin(s) iscontinuously added to, and contacted with, a catalyst in a hydrocarbondiluent. The monomer(s) polymerise to form a slurry of solid particulatepolymer suspended in the polymerisation medium or diluent.

Typically, in the slurry polymerisation process of polyethylene, theslurry in the reactor will comprise the particulate polymer, thehydrocarbon diluent(s), (co) monomer(s), catalyst, chain terminatorssuch as hydrogen and other reactor additives In particular the slurrywill comprise 20-75, preferably 30-70 weight percent based on the totalweight of the slurry of particulate polymer and 80-25, preferably 70-30weight percent based on the total weight of the slurry of suspendingmedium, where the suspending medium is the sum of all the fluidcomponents in the reactor and will comprise the diluent, olefin monomerand any additives; the diluent can be an inert diluent or it can be areactive diluent in particular a liquid olefin monomer; where theprincipal diluent is an inert diluent the olefin monomer will typicallycomprise 2-20, preferably 4-10 weight percent of the slurry.

The slurry is pumped around the relatively smooth path-endless loopreaction system at fluid velocities sufficient (i) to maintain thepolymer in suspension in the slurry and (ii) to maintain acceptablecross-sectional concentration and solids loading gradients.

It has now been found that, for high solids loadings, cross-sectionalslurry concentration distributions (as evidenced by fouling, flowvariations and/or heat transfer) can be maintained within acceptableoperating limits whilst increasing the internal diameter of the tubularreactor above that which is conventionally regarded as operationallyreliable. This is contrary to what the man skilled in the art wouldbelieve to be the case in the light of conventional process conditionswhere the internal diameter of the reactor is no greater than 600millimeters and is typically about 500 millimeters.

The solids concentration in the slurry in the reactor will typically beabove 20 vol %, preferably about 30 volume %, for example 20-40 volume%, preferably 25-35 volume % where volume % is [(total volume of theslurry—volume of the suspending medium)/(total volume of theslurry)]×100. The solids concentration measured as weight percentagewhich is equivalent to that measured as volume percentage will varyaccording to the polymer produced but more particularly according to thediluent used. Where the polymer produced is polyethylene and the diluentis an alkane, for example isobutane it is preferred that the solidsconcentration is above 30 in particular above 40 weight % for example inthe range 40-60 preferably 45%-55 weight % based on the total weight ofthe slurry.

It is a feature of the present invention that operation of the processcan be carried out in larger diameter reactors than are conventionallyused in slurry polymerisation without any significant problemsparticularly from fouling at the reactor walls. For example, reactorshaving internal diameters over 700 millimeters, in particular over 750for example over 850, preferably between 700 and 800 millimetres can beused where historically there would have been increased concern. It ispreferred that greater than 50%, in particular greater than 70%, forexample greater than 85% of the total length of the reactor has aninternal diameter over 700 millimeters, in particular over 750millimeters, for example between 700 and 800 millimeters A particularadvantage of this invention is therefore that high slurry concentrationsat relatively low circulation velocities and relatively high reactorloop diameters can be used. A further embodiment of the presentinvention is a process comprising polymerising in a loop reactor anolefin monomer optionally together with an olefin comonomer in thepresence of a polymerisation catalyst in a diluent to produce a slurrycomprising solid particulate olefin polymer and the diluent wherein theFroude number is maintained at or below 20, preferably 3-10 and theinternal diameter of the reactor is in the range 700-800 millimeters.

The Froude number is preferably maintained at or below 30, for examplein the range 20 to 1 preferably in the range 15 to 2, more preferably inthe range 10 to 3. The Froude number is a dimensionless parameterindicative of the balance between the suspension and settling tendenciesof particles in a slurry. It provides a relative measure of the momentumtransfer process to the pipe wall from particles compared to the fluid.Lower values of the Froude number indicate stronger particle-wall(relative to fluid-wall) interactions. The Froude number (Fr) is definedas v²/(g(s−1)D) where v is the average velocity of the slurry, g is thegravitational constant, s is the specific gravity of the solid in thediluent and D is the internal pipe diameter. The specific gravity of thesolid polymer which is the ratio of the density of the polymer to thedensity of the suspending medium is based on the annealed density of thedegassed polymer after being substantially devolatilised and immediatelyprior to any extrusion as measured using method ISO1183A.

It has been found that reactors can be designed and operated at specificpressure drop both per unit reactor length and per mass of polymer andtotal pressure drop for the loop less than that taught as beingrequired, particularly at high solids loadings and large reactordiameters. This invention permits total loop pressure drops of less than1.3 bar, particularly less than 1 bar even for polymer production ratesof above 25, even above 45 tonnes per hour It is possible to employ oneor more than one pump in the loop preferably on one or more horizontalsections; these can be located on the same horizontal section or ondifferent sections. The pump or pumps can be of the same diameter orlarger or smaller diameter preferably of the same diameter as theinternal diameter of the section of the reactor where the pump or pumpsare located. It is preferable to employ a single pump and it is afeature of the present invention that requirements for number and powerof pump(s) is less onerous than for conventional processes.

Reactor size is typically over 20 m³ in particular over 50 m³ forexample 75-200 m³ preferably in the range 100-175 m³

Use of higher internal diameter reactors as set out herein above enablesreactors, for example having volumes of greater than 80 m³, to be builtwith a reactor length to diameter ratio of less than 500, preferablyless than 400 more preferably less than 250. Reduction in reactor lengthto diameter ratio minimises compositional gradients around the reactionloop and enables production rates of greater than 25 tonnes (perreactor) per hour to be achieved with only a single point ofintroduction for each reagent around the reaction loop. Alternatively itis possible to have multiple inlets into the loop reactor for reactants(e.g. olefins), catalyst, or other additives.

The pressure employed in the loop will be sufficient to maintain thereaction system ‘liquid full’ i.e. there is substantially no gas phase.Typical pressures used are between 1-100 bara, preferably between 30 to50 bara. In ethylene polymerization the ethylene partial pressure willtypically be in the range 0.1 to 5 MPa, preferably from 0.2 to 2 MPa,more particularly from 0.4 to 1.5 MPa. The temperatures selected aresuch that substantially all of the polymer produced is essentially (i)in a non-tacky and non-agglomerative solid particular form and (ii)insoluble in the diluent. The polymerization temperature depends on thehydrocarbon diluent chosen and the polymer being produced. In ethylenepolymerisation the temperature is generally below 130 C, typicallybetween 50 and 125 C., preferably between 75 and 115 C. For example inethylene polymerisation in isobutane diluent, the pressure employed inthe loop is preferably in the range 30-50 bara, the ethylene partialpressure is preferably in the range 0.2-2 MPa and the polymerisationtemperature is in the range 75-115 C. The space time yield which isproduction rate of polymer per unit of loop reactor volume for theprocess of the present invention is in the range 0.1-0.4 preferably0.2-0.35 ton/hour/m³.

The process according to the invention applies to the preparation ofcompositions containing olefin (preferably ethylene) polymers which cancomprise one or a number of olefin homo-polymers and/or one or a numberof copolymers. It is particularly suited to the manufacture of ethylenepolymers and propylene polymers. Ethylene copolymers typically comprisean alpha-olefin in a variable amount which can reach 12% by weight,preferably from 0.5 to 6% by weight, for example approximately 1% byweight.

The alpha mono-olefin monomers generally employed in such reactions areone or more 1-olefins having up to 8 carbon atoms per molecule and nobranching nearer the double bond than the 4-position. Typical examplesinclude ethylene, propylene, butene-1, pentene-1, hexene-1 and octene-1,and mixtures such as ethylene and butene-1 or ethylene and hexene-1.Butene-1, pentene-1 and hexene-1 are particularly preferred comonomersfor ethylene copolymerisation.

Typical diluents employed in such reactions include hydrocarbons having2 to 12, preferably 3 to 8, carbon atoms per molecule, for examplelinear alkanes such as propane, n-butane, n-hexane and n-heptane, orbranched alkanes such as isobutane, isopentane, isooctane and2,2,-dimethylpropane, or cycloalkanes such as cyclopentane andcyclohexane or their mixtures. In the case of ethylene polymerization,the diluent is generally inert with respect to the catalyst, cocatalystand polymer produced (such as liquid aliphatic, cycloaliphatic andaromatic hydrocarbons), at a temperature such that at least 50%(preferably at least 70%) of the polymer formed is insoluble therein.Isobutane is particularly preferred as the diluent for ethylenepolymerisation.

The operating conditions can also be such that the monomers (e.g.ethylene, propylene) act as the diluent as is the case in so called bulkpolymerisation processes. The slurry concentration limits in volumepercent have been found to be able to be applied independently ofmolecular weight of the diluent and whether the diluent is inert orreactive, liquid or supercritical. Propylene monomer is particularlypreferred as the diluent for propylene polymerisation

Methods of molecular weight regulation are known in the art. When usingZiegler-Natta, metallocene and tridentate late transition metal typecatalysts, hydrogen is preferably used, a higher hydrogen pressureresulting in a lower average molecular weight. When using chromium typecatalysts, polymerization temperature is preferably used to regulatemolecular weight.

In commercial plants, the particulate polymer is separated from thediluent in a manner such that the diluent is not exposed tocontamination so as to permit recycle of the diluent to thepolymerization zone with minimal if any purification. Separating theparticulate polymer produced by the process of the present inventionfrom the diluent typically can be by any method known in the art forexample it can involve either (i) the use of discontinuous verticalsettling legs such that the flow of slurry across the opening thereofprovides a zone where the polymer particles can settle to some extentfrom the diluent or (ii) continuous product withdrawal. via a single ormultiple withdrawal ports, the location of which can be anywhere on theloop reactor but is preferably adjacent to the downstream end of ahorizontal section of the loop. Any continuous withdrawal ports willtypically have an internal diameter in the range 2-25, preferably 4-15,especially 5-10 cm This invention permits large scale polymerisationreactors to be operated with low diluent recovery requirements. Theoperation of large diameter reactors with high solids concentrations inthe slurry minimises the quantity of the principal diluent withdrawnfrom the polymerisation loop. Use of concentrating devices on thewithdrawn polymer slurry, preferably hydrocylones (single or in the caseof multiple hydrocyclones in parallel or series), further enhances therecovery of diluent in an energy efficient manner since significantpressure reduction and vaporisation of recovered diluent is avoided.

It has been found that the slurry concentration in the reactor loop canbe optimised by controlling the average particle size and/or theparticle size distribution of the powder within the reactor loop. Theprincipal determinant of the average particle size of the powder is theresidence time in the reactor. The particle size distribution of thecatalyst can be affected by many factors including the particle sizedistribution of the catalyst fed to the reactor, the initial and averagecatalyst activity, the robustness of the catalyst support andsusceptibility of the powder to fragment under reaction conditions.Solids separating devices (such as hydrocyclones) can be used on theslurry withdrawn from the reactor loop to further assist in control ofthe average particle size and the particle size distribution of thepowder in the reactor. The location of the withdrawal point for theconcentrating device and the design and operating conditions of theconcentrating device system, preferably the at least one hydrocyclonerecycle loop, also enables the particle size and particle sizedistribution within the reactor to be controlled. The average particlesize is preferably between 100 and 1500 microns, most preferably between250 and 1000 microns.

The withdrawn, and preferably concentrated, polymer slurry isdepressurised, and optionally heated, prior to introduction into aprimary flash vessel. The stream is preferably heated afterdepressurisation.

The diluent and any monomer vapors recovered in the primary flash vesselare typically condensed, preferably without recompression and reused inthe polymerization process. The pressure of the primary flash vessel ispreferably controlled to enable condensation with a readily availablecooling medium (e.g. cooling water) of essentially all of the flashvapour prior to any recompression. typically such pressure in saidprimary flash vessel will be 4-25 for example 10-20, preferably 15-17bara. The solids recovered from the primary flash vessel is preferablypassed to a secondary flash vessel to remove residual volatiles.Alternatively the slurry may be passed to a flash vessel of lowerpressure than in the above mentioned primary vessel such thatrecompression is needed to condense the recovered diluent. Use of a highpressure flash vessel is preferred.

The process according to the invention can be used to produce resinswhich exhibit specific density in the range 0.890 to 0.930 (lowdensity), 0.930 to 0.940 (medium density) or 0.940 to 0.970 (highdensity).

The process according to the invention is relevant to all olefinpolymerisation catalyst systems, particularly those chosen from theZiegler-type catalysts, in particular those derived from titanium,zirconium or vanadium and from thermally activated silica or inorganicsupported chromium oxide catalysts and from metallocene-type catalysts,metallocene being a cyclopentadienyl derivative of a transition metal,in particular of titanium or zirconium.

Non-limiting examples of Ziegler-type catalysts are the compoundscomprising a transition metal chosen from groups IIIB, IVB, VB or VIB ofthe periodic table, magnesium and a halogen obtained by mixing amagnesium compound with a compound of the transition metal and ahalogenated compound. The halogen can optionally form an integral partof the magnesium compound or of the transition metal compound.

Metallocene-type catalysts may be metallocenes activated by either analumoxane or by an ionising agent as described, for example, in PatentApplication EP-500,944-A1 (Mitsui Toatsu Chemicals).

Ziegler-type catalysts are most preferred. Among these, particularexamples include at least one transition metal chosen from groups IIIB,IVB, VB and VIB, magnesium and at least one halogen. Good results areobtained, with those comprising:

from 10 to 30% by weight of transition metal, preferably from 15 to 20%by weight,

from 20 to 60% by weight of halogen, preferably from 30 to 50% by weight

from 0.5 to 20% by weight of magnesium, usually from 1 to 10% by weight,

from 0.1 to 10% by weight of aluminium, generally from 0.5 to 5% byweight,

the balance generally consists of elements arising from the productsused for their manufacture, such as carbon, hydrogen and oxygen. Thetransition metal and the halogen are preferably titanium and chlorine.

Polymerisations, particularly Ziegler catalysed ones, are typicallycarried out in the presence of a cocatalyst. It is possible to use anycocatalyst known in the art, especially compounds comprising at leastone aluminium-carbon chemical bond, such as optionally halogenatedorganoaluminium compounds, which can comprise oxygen or an element fromgroup I of the periodic table, and aluminoxanes. Particular exampleswould be organoaluminium compounds, of trialkylaluminiums such astriethylaluminium, trialkenylaluminiums such as triisopropenylaluminium,aluminium mono- and dialkoxides such as diethylaluminium ethoxide, mono-and dihalogenated alkylaluminiums such as diethylaluminium chloride,alkylaluminium mono- and dihydrides such as dibutylaluminium hydride andorganoaluminium compounds comprising lithium such as LiAl(C₂H₅)₄.Organoaluminium compounds, especially those which are not halogenated,are well suited. Triethylaluminium and triisobutylaluminium areespecially advantageous.

The chromium-based catalyst is preferred to comprise a supportedchromium oxide catalyst having a titania-containing support, for examplea composite silica and titania support. A particularly preferredchromium-based catalyst may comprise from 0.5 to 5 wt % chromium,preferably around 1 wt % chromium, such as 0.9 wt % chromium based onthe weight of the chromium-containing catalyst. The support comprises atleast 2 wt % titanium, preferably around 2 to 3 wt % titanium, morepreferably around 2.3 wt % titanium based on the weight of the chromiumcontaining catalyst. The chromium-based catalyst may have a specificsurface area of from 200 to 700 m.sup.2/g, preferably from 400 to 550m.sup.2/g and a volume porosity of greater than 2 cc/g preferably from 2to 3 cc/g.

Silica supported chromium catalysts are typically subjected to aninitial activation step in air at an elevated activation temperature.The activation temperature preferably ranges from 500 to 850.degree. C.,more preferably 600 to 750.degree. C.

The reactor loop can be used to make monomodal or multimodal, forexample bimodal, polymers. The multi-modal polymers can be made in asingle reactor or in multiple reactors. The reactor system can compriseone or more loop reactors connected in series or in parallel. Thereactor loop may also be preceded or followed by a polymerisationreactor that is not a loop reactor.

In the case of series reactors, a first reactor of the series issupplied with catalyst and the cocatalyst in addition to the diluent andmonomer, and each subsequent reactor is supplied with, at least,monomer, in particular ethylene and with the slurry arising from apreceding reactor of the series, this mixture comprising the catalyst,the cocatalyst and a mixture of the polymers produced in a precedingreactor of the series. It is optionally possible to supply a secondreactor and/or, if appropriate, at least one of the following reactorswith fresh catalyst and/or cocatalyst. However, it is preferable tointroduce the catalyst and the cocatalyst exclusively into a firstreactor.

In the case where the plant comprises at least two reactors in series,the polymer of highest melt index and the polymer of lowest melt indexcan be produced in two adjacent or non-adjacent reactors in the series.Hydrogen is maintained at (i) a low (or zero) concentration in thereactor(s) manufacturing the high molecular weight components, e.g.hydrogen percentages including between 0-0.1 vol % and at (ii) a veryhigh concentration in the reactor(s) manufacturing the low molecularweight components e.g. hydrogen percentages between 0.5-2.4 vol %. Thereactors can equally be operated to produce essentially the same polymermelt index in successive reactors.

Particular sensitivity to operating in large diameter reactors (andassociated cross-sectional compositional, thermal or particulategradients) has however been related to production of polymer resinswhere polymer of either high or low molecular weight resins has beenknown to lead to increased fouling concerns. Particularly when producingpolymers of molecular weights less than 50 kDaltons or greater than 150kDaltons. These concerns have particularly been confirmed to beaccentuated at low polymer solids concentrations in the reactor loop.When producing polymers of molecular weights less than 50 kDaltons orgreater than 200 kDa (or melt index below 0.1 and above 50) in largediameter reactors it has however surprisingly been discovered thatfouling is decreased when solids loadings are increased to above 20 vol%, particularly above 30 vol %. The present invention further comprisesa loop reactor of a continuous tubular construction having at least 2horizontal sections and at least 2 vertical sections wherein theinternal diameter of at least 50% of the total length of the reactor isat least 700 millimeters

EXAMPLE

In an elongated closed loop tubular reactor having an internal diameterof 765 millimetres and a volumetric capacity of 167.5 m³, ethylene iscopolymerised with hexene-1 at a temperature of 93° C. and a pressure of41 bara in isobutane as diluent and using a Chromium catalyst to producea copolymer having a MI5 of 0.85 g/10 Minutes and a density of 938kg/m³. An essentially constant solids loading of about 56 wt % ismaintained for a period of several days. The reactor circulation pumppower as measured by the power transducer on the pump motor controlsystem and heat transfer coefficient as measured by monitoring coolantwater flow and coolant water temperature change compared to reactortemperature remain stable, resulting in essentially no change in eitherparameter respectively, indicating that there is no detectable foulingof the reactor as evidenced by a build up of polymer on the walls of thereactor, and that flow is stable and well distributed as evidenced bythe stable pump power readings

1. A process comprising polymerising in a loop reactor of continuoustubular construction an olefin monomer optionally together with anolefin comonomer in the presence of a polymerisation catalyst in adiluent to produce a slurry comprising solid particulate olefin polymerand the diluent wherein the internal diameter of at least 50% of thetotal length of the reactor is at least 700 millimeters, wherein thesolids concentration in the reactor is at least 20 volume % and whereinthe Froude number is maintained below
 30. 2. A process as claimed inclaim 1 wherein the solids concentration in the reactor is in the range25-35 volume %.
 3. A process as claimed in either claim 1 wherein theinternal diameter of at least 50% of the total length of the reactor isover 750 millimeters.
 4. A process as claimed in claim 1 wherein theinternal diameter of at least 70% of the total length of the reactor isover 700 millimeters.
 5. A process as claimed in claim 4 wherein theinternal diameter of at least 85% of the total length of the reactor isover 700 millimeters.
 6. A process as claimed in claim 1 wherein thespace time yield is in the range of 0.2 to 0.35 ton/hour/m³.
 7. A loopreactor of a continuous tubular construction having at least 2horizontal sections and at least 2 vertical sections wherein theinternal diameter of at least 50% of the total length of the reactor isat least 700 millimeters, wherein the ratio of reactor length tointernal diameter is less than
 250. 8. A loop reactor as claimed inclaim 7 wherein the internal diameter of at least 50% of the totallength of the reactor is over 750 millimeters.
 9. A loop reactor asclaimed in claim 7 wherein the internal diameter of at least 70% of thetotal length of the reactor is at least 700 millimeters.
 10. A loopreactor as claimed in claim 7 wherein the internal diameter of at least85% of the total length of the reactor is at least 700 millimeters. 11.A loop reactor as claimed in claim 7 wherein the reactor size is over20m³.
 12. A loop reactor as claimed in claim 11 wherein the reactor sizeis in the range 100-175m³.
 13. A process as claimed in claim 1 whereinthe polymer produced has a molecular weight of less than 50 k Daltons.14. A process as claimed in claim 1 wherein the polymer produced has amolecular weight of greater than 150 k Daltons.